Integrated direct coal liquefaction process and system

ABSTRACT

An integrated direct coal liquefaction and upgrading process and system in which feed coal is mixed with a coal liquefaction nominal 650° F.+ (343° C.+) fraction from atmospheric fractionator, vacuum fractionator bottoms, and recycled catalyst containing vacuum gas oil (VGO) from and upgrader and pumped to the input of the liquefaction reactor where it is mixed with preheated hydrogen. Part of the VGO fraction from the liquefaction vacuum fractionator is mixed with make-up catalyst and catalyst containing VGO fraction from the upgrader atmospheric fractionator and fed with preheated hydrogen to the input of the upgrading reactors. The output of the upgrading reactors is cooled in hot and cold separators to recover hydrogen containing gaseous components and the liquid fraction is separated in the upgrader atmospheric fractionator into product and a catalyst containing unconverted VGO stream, a portion of which forms the recycled VGO to the liquefaction slurry mix tank. A portion of the catalyst containing vacuum fractionator bottoms are sent as fuel to a cement plant in which the catalyst is gasified and recovered.

FIELD OF THE INVENTION

The present invention relates to integrated direct coal liquefaction and upgrading processes and systems in which catalyst is circulated between liquefaction and upgrading reactors and catalyst entrained in reactor bottoms are recovered in a cement plant.

BACKGROUND

Prior direct coal liquefaction processes have been proposed involving one or more of dispersed catalysts, slurry reactors, bottoms recycle, integration of liquefaction with certain types of upgrading, and sending liquefaction bottoms to a cement plant.

In U.S. Pat. No. 7,763,167 B2, a process for direct coal liquefaction is described which combines suspended bed reactors for liquefying coal with a separate suspended bed upgrader for both upgrading the product and producing a hydrogen donor solvent. The donor solvent is used for slurrying the feed coal prior to pumping the slurry up to liquefaction pressure. The donor solvent stream is free of ash and catalyst. The catalyst addition rate to liquefaction is 0.5 to 1.0 wt % on dry coal. Catalyst is utilized once through and is disposed of in the residue from the vacuum tower.

In U.S. Pat. No. 4,602,992, a process for direct liquefaction is described that integrates coal liquefaction with an integrated refining stage. After liquefaction in a slurry phase reactor, the high boiling fractions are separated from the lower boiling vapor fractions, and the lower boiling fraction is subsequently hydrogenated in a fixed bed reactor. Part of the hydrogenated stream, which is free of ash and catalyst, is removed as product and part is recycled for slurrying the fresh coal. Catalyst for liquefaction is utilized once through and is disposed of in the residue from the vacuum tower.

In U.S. 2011/0042272 A1, a process for integrated direct coal liquefaction and hydrotreating is described. Here again, vapors from the hot separator are sent directly to a product hydrotreating step, in this case an ebullated bed reactor. Liquefaction occurs in two ebullated bed reactors operating in series. A portion of the supported catalyst is cascaded from the hydrotreater to the first liquefaction reactor and subsequently to the second reactor. The coal is slurried with a catalyst free bottoms recycle stream boiling above 500° F. (260° C.).

In WO 2014/110085 A1, a direct coal liquefaction process is described for liquefaction of high inertinite content coals. The process uses a dispersed catalyst and uses atmospheric fractionator bottoms, vacuum fractionator bottoms, and a process derived 700/1000° F. fraction referred to as vacuum gas oil (VGO) as a coal solvent. Sending bottoms to a cement plant is identified as one of the options for bottoms disposal.

In U.S. Pat. No. 4,249,951, U.S. Pat. No. 4,260,421, and EP 1985 596 A1 processes are described for using mineral-containing residue from coal liquefaction processes to produce a hydraulic cement. The residue is mixed with siliceous and calcareous materials and compacted. The compacted solids are introduced into a kiln and fired to produce hydraulic cement clinker. The clinker is removed from the kiln and ground to produce cement.

SUMMARY OF THE INVENTION

In accordance with the invention, there is provided an integrated direct coal liquefaction and upgrading process and system in which a slurry of a nominal 650° F.+ (343° C.+) fraction from a atmospheric fractionator mixed with feed coal containing solid carbonaceous material, molybdenum catalyst containing vacuum gas oil (VGO) and vacuum fractionator bottoms is mixed with preheated hydrogen and fed to the input of a direct coal liquefaction (DCL) slurry reactor. The product of the DCL reactor is separated into components including a hydrogen containing gaseous stream, a C3/650° F. product stream, and a VGO stream and catalyst containing vacuum fractionator bottoms, a portion of said vacuum fractionator bottoms being used in the formation of said slurry. A molybdenum containing catalyst or catalyst precursor is added to DCL reactor produced VGO and the mixture, along with preheated hydrogen, is fed to the input of an upgrading reactor. The product of the upgrading reactor is separated into components including a liquid product stream and a molybdenum catalyst containing unconverted VGO stream. A portion of the catalyst containing unconverted VGO is added to the DCL reactor produced VGO being fed to the input of the upgrading reactor along with molybdenum containing catalyst or catalyst precursor. Another portion of the catalyst containing upgrading reactor produced VGO is used as the catalyst containing VGO constituent of the slurry being fed to the DCL reactor.

In accordance with a very important aspect of the invention, because the amount of VGO from liquefaction sent to the upgrading reactor is only about 15% of the feed coal to supplied to the DCL reactor on a DAF basis, the catalyst concentration in the upgrading reactor is approximately six fold higher than in the DCL reactor. Additionally, the concentration of catalyst in the DCL reactor is increased because of the recycle of the fractionator bottoms to the DCL reactor. In accordance with another aspect of the invention, a portion of said vacuum fractionator bottoms is fed as fuel to a cement plant, and molybdenum containing catalyst is recovered in the cement plant from the vacuum fractionator bottoms.

In accordance with another aspect of the invention hot gases produced by the DCL slurry reactor are used in a heat exchanger to preheat the hydrogen being fed to the input of the DCL reactor and the slurry is fed to the input of the DCL reactor without being preheated other than by the preheated hydrogen.

In accordance with a still further aspect of the invention, if necessary, the preheated hydrogen is further heated in a hydrogen furnace before being mixed with the slurry.

Other important features of the process and system of the invention include the facts that

a. The makeup of the dispersed catalyst is minimized by the recovery of catalyst in the cement plant and the recirculation to the DCL reactor of catalyst entrained in the bottoms stream from the atmospheric and vacuum fractionators. b. The low severity direct coal liquefaction process produces a product slate that is rich in vacuum gas oil (VGO) and lean in C1-C4 production, and the high catalyst concentration upgrading process converts VGO into liquid products with no ash present in the upgrading reactors resulting in a minimum of C1-C4 production. c. Ash present in the feed coal ends up in the cement product and thus eliminates disposal costs for the waste ash stream. d. DCL reactor produced bottoms are used as fuel in the cement plant rather than coal, thus avoiding additional CO₂ production from the cement plant and incinerating hydrocarbons present in the DCL reactor produced bottoms. e. CO₂ from the direct coal liquefaction and upgrading processes is minimized by maximum use of the heat of reaction via recycle of hot unconverted product back to the mix tanks for the direct coal liquefaction and upgrading processes and via use of a very active dispersed catalyst for both processes that allows initiation of both direct coal liquefaction and upgrading, in slurry reactors, at a low inlet temperature. f. Operation with higher solids levels in the slurry DCL reactor results in lower gas hold-up in the reactor. The high solids levels are the result of recycle of ash and catalyst in the bottoms streams. g. Thermal efficiency exceeds 70% for a plant that utilizes coal gasification for production of hydrogen and over 75% for a plant producing hydrogen from natural gas. The high thermal efficiency is the result of maximum use of the heat of reaction in both the liquefaction unit and the upgrader and the use of bottoms for fuel in the cement plant rather than additional coal. h. Economics are improved because of lower catalyst cost, improved product selectivity, lower fuel requirement, and lower investment. Investment is reduced because of elimination of slurry heat exchangers, slurry preheat furnaces, ash disposal, and elimination of bottoms processes such as gasification or power generation.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic diagram of a direct coal liquefaction system suitable for use in the illustrated embodiment of the invention

FIG. 2 is a schematic diagram of a preferred embodiment of upgrading system of the invention.

FIG. 3 is a diagram of a typical temperature profile for a cement plant.

FIG. 4 is a schematic diagram of a system for recovering MoO3 from a cement kiln.

DESCRIPTION OF THE PREFERRED EMBODIMENTS

The DCL process of this invention achieves high coal conversion and liquid yields without the need for either slurry heat exchangers or slurry preheat furnace. All heat input to the liquefaction reactor is achieved by recycle of hot atmospheric and vacuum pipestill bottoms and a hydrogen treat gas heat exchanger and, if necessary, a hydrogen preheat furnace.

An embodiment of a reactor system for performing the direct coal liquefaction in accordance with the invention is shown in FIG. 1. A coal containing solid carbonaceous material feed is dried and crushed in a conventional gas swept roller mill (not shown) to a moisture content of 1 to 4%. The crushed and dried coal containing solid carbonaceous material 101 and a catalyst containing 700° F.+ VGO stream 103 from the upgrader of FIG. 2 are fed into a slurry mixing tank 105 where it are mixed with a 600 to 700° F.+ (316 to 371° C.+)(nominally 650° F.+) fraction 107 from the atmospheric pipestill (APS) 111 and a portion of the bottoms 133 from the vacuum pipestill (VPS) 113 (also referred to more generally as atmospheric and vacuum fractionators, respectively) and with a nominal 650° F. to 1000° F. (343° C. to 538° C.) VGO fraction from line 131 to form a slurry stream.

The slurry mix tank operating temperature is set by controlling the relative amounts of the nominal 650° F.+ fraction 107 from the APS 111, VPS bottoms, and VGO being fed thereto. Typical operating temperature ranges from 500 to 700° F. (260 to 371° C.) and more preferably about 600° F. (316° C.). From the slurry mix tank 105, the catalyst containing slurry is delivered to the slurry pump 115.

In the illustrated embodiment, the slurry leaves the mixing tank 105 at about 600° F. (316° C.). Most of the moisture in the coal is driven off in the mixing tank 105 due to the hot bottoms being fed thereto. Such moisture and volatiles are sent to separation. The coal in the slurry leaving the slurry mixing tank 105 has about 0.1 to 1.0% moisture.

The coal slurry is pumped from the mixing tank 105 and the pressure raised to about 2,000 to 3,000 psig (138 to 206 kg/cm² g) by the slurry pumping system 115. The resulting high pressure slurry is mixed with preheated hydrogen rich treat gas. The hydrogen treat gas is preheated in heat exchanger 117 and, if necessary, in preheat furnace 119. Heat for the hydrogen exchanger comes from the overhead from hot separator 121. Heat exchanger 117 can be an air or water cooled exchanger.

The coal slurry and hydrogen mixture is fed to the input 123 of the first stage of the series-connected liquefaction reactors 125 at about 660 to 700° F. (349 to 371° C.) and 2,000 to 3,000 psig (138 to 206 kg/cm² g). The reactors 125 are up-flow tubular vessels, the total length of the three reactors being 50 to 250 feet. The temperature rises from one reactor stage to the next as a result of the highly exothermic coal liquefaction reactions. In order to maintain the maximum temperature in each stage below about 800 to 900° F. (427 to 482° C.), additional hydrogen treat gas is preferably injected between reactor stages. The hydrogen partial pressure in each stage is preferably maintained at a minimum of about 1,000 to 2,000 psig (69 to 138 kg/cm² g).

The effluent from the last stage of liquefaction reactor 125 is fed to the hot separator 121 in which it is separated into a gas stream and a liquid/solid stream. The gas stream is sent to the heat exchanger 117 in which it serves, optionally together with the hydrogen furnace 119, to preheat the hydrogen being fed to the liquefaction reactor input 123. The liquid/solid stream from the hot separator 121 is let down in pressure and fed to the APS 111. After passing through the heat exchanger 117, the gas stream from the hot separator 121 is cooled in heat exchanger 127 and fed to the cold separator 129 to condense out the liquid vapors of naphtha, distillate, and solvent and processed to remove H₂S and CO₂.

Most of the remaining processed gas from the cold separator 129 is then sent to the hydrogen recovery system 131 for further processing by conventional means to recover the hydrogen contained therein, which is then recycled via the heat exchanger 117 and the hydrogen furnace 119 to be mixed with the coal slurry. The remaining portion of the processed gas is purged to prevent buildup of light ends in the recycle loop. Hydrogen recovered therefrom can be used in the downstream hydro-processing upgrading system. Make-up hydrogen is added on line 122 to maintain an adequate hydrogen partial pressure in the liquefaction reactors.

The depressurized liquid/solid stream and the hydrocarbons condensed during the gas cooling are sent to the APS 111 where they are separated into light ends, naphtha, distillate and bottoms fractions. The light ends are processed to recover hydrogen and C₁-C₄ hydrocarbons that can be used for fuel gas and other purposes. The naphtha is hydrotreated to saturate diolefins and other reactive hydrocarbon compounds. The 160° F.+ fraction of the naphtha can be hydrotreated and catalytically reformed to produce gasoline. The distillate fraction can be hydrotreated to produce products such as diesel and jet fuel. A portion of the 600 to 700° F.+ (316 to 371° C.+)(nominally, 650° F.+ (343° C.+)) fraction is recycled to the slurry mix tank 105 on line 107. The 600 to 700° F.− (316 to 371° C.−)(nominally, 650° F.− (343° C.+)) light ends, naphtha and distillate fractions are taken off the APS 111 on lines indicated schematically as line 112.

The remaining nominal 650° F.+ (343° C.+) fraction produced from the atmospheric fractionator 111 is fed to the VPS 113 wherein it is separated into a nominal 650 to 950/1100° F. (343 to 510/593° C.) VGO fraction and a 950/1100° F.+ (510/593° C.+) bottoms fraction (nominally 650 to 1000° F. (343 to 538° C.) and 1000° F.+ (538° C.+) fractions). A portion of the VGO fraction on line 131 is added to the nominal 650° F.+ (343° C.+) stream from the APS being recycled to the slurry mix tank 105. The APS 111 is preferably operated at a high enough pressure so that a portion of the nominal 650° F.+ (343° C.+) fraction can be recycled to the slurry mixing tank 105 without pumping.

A portion of the 1000° F.+ fraction from the VPS 113 is sent via line 133 to be used as fuel in the cement plant illustrated in FIG. 4, as described below. The remainder of the 1000° F.+ fraction is recirculated to the slurry mix tank 105. A portion of the VGO produced by the VPS 113, is sent via line 131 to the upgrading system illustrated in FIG. 2 of the drawings with the remainder of the VGO being recycled to the slurry mix tank 105. Substantially all of the catalyst contained in the effluent from the liquefaction reactors 125 ends up in the bottoms from the atmospheric fractionator 111 and the vacuum fractionator 113 from which it is fed either to the slurry mix tank 105 and recirculated to the liquefaction reactors 125, or is sent to the cement plant (FIG. 4) via line 133. The VGO fed from the vacuum fractionator 113 sent to the upgrading system (FIG. 2) via line 131 contains substantially no catalyst. Additional hydrogen for the process can also be produced via steam reforming of natural gas or via gasification of coal. Catalysts useful in DCL processes also include those disclosed in U.S. Pat. Nos. 4,077,867, 4,196,072 and 4,561,964, the disclosures of which are hereby incorporated by reference in their entirety.

FIG. 2 of the drawings illustrates a preferred embodiment of the upgrading portion of the integrated direct coal liquefaction/upgrading method and system of the invention. The VGO stream from line 131 (FIG. 1) is fed to the mix tank 201 where is mixed with dispersed molybdenum catalyst or catalyst precursor from line 203 and unconverted VGO bottoms from the atmospheric fractionator 205. Typical operating temperature of the mix tank 201 is 550 to 750° F. and preferably 600 to 700° F. The catalyst in the illustrated embodiment is preferably in the form of a 2-10% aqueous water solution of phosphomolybdic acid (PMA) in an amount that is equivalent to adding between 50 wppm and 2% molybdenum relative to the dry coal feed.

The catalyst/VGO output from the mix tank 201 is pumped and the pressure increased to about 2,000 to 3,000 psig by the pumping system 207. The resulting high pressure, hot stream 209 is mixed with preheated hydrogen 211 and fed to the input 213 of the upgrading reactors 215. The upgrader reactors 215 are up-flow, tubular reactors. At least two, and preferably three or more reactors are used in series; thus approaching a plug flow reactor. Hydrogen quench is required between reactors to enable control of the temperature profile.

The hydrogen rich stream 211 is preheated in heat exchanger 217 and, if necessary, in hydrogen furnace 219. Depending upon the hydrogen recycle rate 221, the hydrogen furnace may or may not be required. The heat for the heat exchanger 217 comes from the overhead from the hot separator 223. The temperature of the hot overhead stream is reduced in the heat exchanger 217 as a result of the heat exchange with the hydrogen rich stream 221. Stream 225 is further cooled via air or water in heat exchanger 227 and sent to the cold separator 229, in which it is scrubbed to remove H₂S and ammonia. A portion 231 of the hydrogen rich stream from the cold separator 229 is recycled to back to the upgrading reactor 215, and a purge stream 233 is sent to hydrogen recovery. The recycle gas 231 is compressed in the recycle gas compressor 235 to offset the pressure loss in the system. Make-up hydrogen 237 is added to offset hydrogen consumption in the system and purge gas 233.

Liquid streams from the hot separator 223 and cold separator 229 are sent to the atmospheric fractionator 205. The output from the atmospheric fractionator 205 boiling below a cut point of 650 to 700° F. (nominally, 700° F.−) is removed as product on line 234 and can be combined with nominal 650° F.− product from the coal liquefaction unit. A portion of the unconverted catalyst containing 700° F.+ VGO from the atmospheric fractionator 205 is recycled via line 103 to the slurry mix tank 105 (FIG. 1).

This stream contains a high concentration of dispersed catalyst for use in the direct coal liquefaction unit. The remainder of the VGO stream from the atmospheric fractionator 205 is recycled to the mix tank 201 where it is mixed with catalyst free VGO from the vacuum fractionator 113 (FIG. 1) and make up catalyst.

The quantity of catalyst fed to the mix tank 201 on line 203 is preferably substantially the same as would be fed to the liquefaction slurry reactors 125 (FIG. 1) in a prior art DCL slurry reactor system of the same capacity. Since, however, the catalyst free VGO sent from the vacuum pipestill 113 to the mix tank 201 via line 131 is typically equal to approximately 13 wt % of the feed coal on a DAF basis, the catalyst concentration in the upgrader reactors 215 will be over eight times the catalyst concentration in the coal liquefaction unit (FIG. 1). For example, if 300 wppm of dispersed catalyst on a DAF feed coal basis were added to the mix tank 201, the concentration of dispersed catalyst in the upgrader reactors 215 would be approximately 2400 wppm on the basis of the VGO in the upgrader reactors 215.

Additionally, the concentration of catalyst in the slurry stream sent the liquefaction reactors 125 will be higher than in the catalyst make-up feed to the mix tank 201 on line 203 because of the conversion of a portion of the VGO to 700° F.− products by the upgrading reactors 215. The recycling to the liquefaction reactor 125 of catalyst containing bottoms from the atmospheric fractionator 111 and the vacuum fractionator 113 also increases the concentration of catalyst in the liquefaction reactor 125.

The combination of high catalyst concentration and absence of ash in the upgrading reactors 215 results in an upgrading process that converts VGO into liquid products with a minimum of C1-C4 production.

The slurry upgrading unit can also be utilized to upgrade nominal C5/1000° F., 350/1000° F., or 650/1000° F. streams from the liquefaction unit or a combination of one of these streams with residuum, heavy oil, bitumen, or coal tar.

Referring to FIGS. 3 and 4 of the drawings, an illustrative embodiment of a cement kiln 400 in accordance with the invention is divided into two kiln sections 405 and 407 with differing diameters. Both of these kiln sections rotate. The kiln sections 405 and 407 are connected by a stationary seal 409 to prevent hot gases escaping between the kiln sections. Ingredients 401 for cement manufacture are introduced at the left of kiln section 405 and product is withdrawn at the outlet 411 of kiln section 407. Fuel and air feed 403 to the kilns is introduced at the right of kiln section 407. Effluent gases 413 exit from the left of kiln section 405. Thus, the ingredients 401 for cement and the air and fuel 403 are countercurrent flows in the kiln 400. Because of the countercurrent flow, as indicated in FIG. 3, peak temperatures, which are over 1400° C., occur in the middle portion of the kiln 400. These temperatures are sufficient to sinter the carbonates in the feed ingredients 401 to form the cement and to oxidize and vaporize the molybdenum catalyst.

Nominal 1000° F.+ bottoms from the vacuum fractionator 113, containing unconverted coal, ash and catalyst, are sent via line 133 as all or part of the fuel 403 for the cement kiln 400. The carbonaceous portion of the bottoms is consumed as fuel for the cement kiln 400, thus decomposing the carbonates in the feed (CaCO₃ for example) to cement (CaO). The vacuum fractionator bottoms take the place of all or part of the coal normally used as the fuel in a cement kiln. CO₂ is produced from both the decomposition of carbonates present in the feed ingredients 401, and from the combustion of the fuel used to heat and decompose the feed.

Fuel may also be introduced at an additional inlet (not shown) in the stationary seal 409. Depending on material balance and other considerations, a portion of the fuel may also be introduced at the left-hand end of the kiln 400 together with the cement ingredients 401.

The use of bottoms from the vacuum fractionator 113 as fuel for the cement kiln has several advantages. First, the bottoms replace all or part of the coal fuel, so that there is a reduction in CO2 emissions that would otherwise occur from the use of coal as fuel. Secondly, the ash present in the vacuum fractionator bottoms become part of the product cement, thus eliminating ash disposal. Third, the molybdenum contained in the catalyst present in the fuel feed to the cement plant is recovered for reuse. Fourth, sulfur in the bottoms is converted into SO2 and reacts with the calcium in the cement, thereby preventing SO2 emissions.

Because of the temperature profile in the kiln, the molybdenum containing catalyst in the bottoms stream is converted to MoO₃ which vaporizes in the high temperature central section of the kiln sections 405 and 407 (MoO₃ sublimes at 1,155° C.). As the hot vapor flows to the left in the kiln 400, it contacts cold feed and is condensed on the solid feed 401. As the now molybdenum containing feed 401 flows to the right in the kiln 400, the MoO₃ is again vaporized and joins the hot gases flowing to the left in the kiln 400. Thus, the MoO₃ becomes trapped in the middle of the kiln. An outlet 415 in the stationary seal 409 between the two kiln sections 405 and 407, allows the removal of a portion of the hot MoO₃ containing gas stream from the kiln 400. This stream is cooled to condense the MoO₃ and dissolved by contact with phosphoric acid to produce phosphomolybdic acid (PMA) catalyst precursor that is used to prepare molybdenum containing catalyst to be sent to the mix tank 201 as a part of the catalyst make-up stream 203.

Example 1

If the desired concentration of dispersed molybdenum catalyst in the direct coal liquefaction reactor 125 is 300 wppm in the above described integrated process, and assuming that the direct coal liquefaction process is operated to produce a synfuel product slate that contains approximately 13 wt % VGO (700/1000° F.) on a DAF coal basis, introduction of the catalyst required for direct coal liquefaction into the upgrader mix tank 201 will result in a molybdenum content of the VGO being fed to the upgrading reactor 215 of 2,400 wppm on a VGO basis.

Converting 90% of the VGO in the upgrader reactor 215 increases the molybdenum concentration in the upgrader atmospheric fractionator bottoms VGO to 2.4 wt %. If the feed to the upgrader mix tank 201 includes a one to one ratio of catalyst free VGO from the liquefaction vacuum fractionator 113 and recycle from the upgrader fractionator 205, the average catalyst concentration in the feed to the upgrading reactors 215 from the mix tank 201 will be 1.3 wt %. At this catalyst loading and a 1/1 recycle to feed ratio, the VGO will be selectively converted in the upgrading reactors 215 to 700° F. minus product with a minimum of light gas production (C1-C4) and heteroatoms, such as sulfur, nitrogen, and oxygen, in the product will be significantly reduced. The VGO stream sent from the upgrading fractionator 205 to the liquefaction slurry mix tank 105 will contain 2.4 wt % catalyst.

After liquefaction, part of the 1000° F.+ bottoms stream from the vacuum fractionator 113 is sent to the cement kiln 400. If 90% of the molybdenum catalyst present in such bottoms is recovered and sent to the mix tank 201 on line 203, the additional required make-up molybdenum catalyst to the mix tank 201 will be 30 wppm or 10 percent of the molybdenum catalyst required in the direct coal liquefaction unit. 

What is claimed is: 1) An integrated direct coal liquefaction and upgrading process, comprising the steps of: a) forming a slurry of feed coal containing solid carbonaceous material mixed with a nominal 650° F.+ (343° C.+) stream, a vacuum gas oil (VGO) stream, molybdenum catalyst containing VGO from an upgrader, and catalyst containing nominal 1000° F.+ (538° C.+) bottoms; b) feeding said slurry and preheated hydrogen to the input of a direct coal liquefaction (DCL) slurry reactor; c) separating the product of said DCL reactor into components including a hydrogen containing gaseous stream, a VGO stream, a nominal 650° F.+ (343° C.+) stream, and catalyst containing nominal 1000° F.+ (538° C.+) bottoms, portions of said nominal 650° F.+(343° C.+) stream, of said VGO stream, and of said bottoms being used as components of said slurry; d) feeding a portion of said DCL reactor product VGO stream, a molybdenum containing catalyst or catalyst precursor, molybdenum catalyst containing VGO and preheated hydrogen to the input of an upgrading reactor; e) separating the product of the upgrading reactor into components including a liquid product stream and a molybdenum catalyst containing VGO stream, said upgrading reactor produced molybdenum catalyst containing VGO stream being used as a component of said slurry and as a component of the feed to the input of the upgrading reactor. 2) The integrated direct coal liquefaction and upgrading process of claim 1 further including the steps of: a) feeding a portion of said nominal 1000° F.+ (538° C.+) bottoms as fuel to a cement plant; and b) recovering molybdenum containing catalyst from said vacuum fractionator bottoms in said cement plant. 3) An integrated apparatus for converting a coal containing solid carbonaceous material feed into hydrocarbon liquids and for upgrading such liquids, comprising: a) a direct coal liquefaction (DCL) reactor for converting such coal containing solid carbonaceous material at elevated temperatures and pressures in the presence of a solvent and a catalyst for producing hydrocarbon products; b) an atmospheric fractionator for separating hydrocarbon products of said DCL reactor into different boiling point fractions, including a nominal 650° F.+ (343° C.+) fraction; c) a vacuum fractionator for separating a portion of said 650° F.+ fraction from the atmospheric fractionator into a nominal 650° F. to 1000° F. (343° C. to 538° C.−) VGO fraction and a catalyst containing nominal 1000° F.+ (538° C.+) bottoms fraction; d) a DCL reactor slurry mix tank for mixing a portion of said 650° F.+ (343° C.+) fraction with feed coal containing solid carbonaceous material, molybdenum catalyst containing VGO, DCL reactor product VGO, and a portion of said catalyst containing bottoms fraction to form a DCL reactor slurry; e) a pump for feeding said DCL reactor slurry to the input of said DCL reactor; f) an upgrading reactor, g) an upgrader slurry mix tank; and h) an upgrader atmospheric fractionator for separating the product of the upgrading reactor into components including a liquid product stream and a molybdenum catalyst containing VGO stream, a portion of the said upgrading reactor produced molybdenum catalyst containing VGO stream being fed to said DCL reactor slurry mix tank and the remainder thereof being fed to said upgrader slurry mix tank, said upgrader slurry mix tank mixing a portion of said DCL reactor product VGO fraction with said molybdenum catalyst containing VGO stream and a molybdenum containing catalyst or catalyst precursor to form an upgrading slurry and feeding said upgrading slurry to the input of said upgrading reactor; and i) heat exchanger means for feeding preheated hydrogen to the inputs of said DCL reactor and said upgrading reactor. 4) The integrated apparatus of claim 3 further including: a) feeding a portion of said nominal 1000° F.+ (538° C.+) bottoms as fuel to a cement plant; and b) recovering molybdenum containing catalyst from said vacuum fractionator bottoms in said cement plant. 5) A cement kiln for recovering molybdenum containing catalyst from coal liquefaction process bottoms comprising: a) first and second longitudinally connected rotatable kiln sections having differing diameters; b) a stationary seal at the connection of said kiln sections for preventing hot gases from escaping between said kiln sections; c) an input at the free end of said first kiln section for introducing Ingredients for cement manufacture into said first kiln section; d) an input at the free end of said second kiln section for introducing molybdenum catalyst containing 1000° F.+ (538° C.+) coal liquefaction process bottoms into said second kiln section; and e) an outlet in said stationary seal for removing a portion of hot MoO₃ containing gas stream from the kiln
 400. 